U.S. patent application number 11/571309 was filed with the patent office on 2007-10-25 for zeolite catalysts.
Invention is credited to Ivar Dahl, Klaus Jens, Trond Myrstad, Ase Slagtern.
Application Number | 20070246400 11/571309 |
Document ID | / |
Family ID | 32800318 |
Filed Date | 2007-10-25 |
United States Patent
Application |
20070246400 |
Kind Code |
A1 |
Jens; Klaus ; et
al. |
October 25, 2007 |
Zeolite Catalysts
Abstract
A process for cracking C.sub.4+ hydrocarbons comprising heating
the hydrocarbons at a temperature of 400-800.degree. C. in the
presence of an H-ZSM-5 catalyst having a silicon to aluminium ratio
of 61-1000.
Inventors: |
Jens; Klaus; (Stathelle,
NO) ; Myrstad; Trond; (Vikhammer, NO) ; Dahl;
Ivar; (Trondheim, NO) ; Slagtern; Ase;
(Trondheim, NO) |
Correspondence
Address: |
GARDNER GROFF GREENWALD & VILLANUEVA. PC
2018 POWERS FERRY ROAD
SUITE 800
ATLANTA
GA
30339
US
|
Family ID: |
32800318 |
Appl. No.: |
11/571309 |
Filed: |
January 5, 2006 |
PCT Filed: |
January 5, 2006 |
PCT NO: |
PCT/EP05/06927 |
371 Date: |
May 5, 2007 |
Current U.S.
Class: |
208/111.01 ;
208/120.01 |
Current CPC
Class: |
B01J 29/40 20130101;
C10G 69/04 20130101; C10G 11/02 20130101 |
Class at
Publication: |
208/111.01 ;
208/120.01 |
International
Class: |
B01J 29/40 20060101
B01J029/40 |
Foreign Application Data
Date |
Code |
Application Number |
Jun 28, 2004 |
GB |
0414442.4 |
Claims
1-21. (canceled)
22. A process for cracking C.sub.4+ hydrocarbon feedstock
comprising saturated cyclic hydrocarbons, said process comprising:
heating the feedstock at a temperature of 400-800.degree. C. in the
presence of an H-ZSM-5 catalyst having a silicon to aluminum ration
of 61-1,000.
23. The process of claim 1, wherein the C4+ hydrocarbon feedstock
comprises at least one C4+ olefin.
24. The process of claim 1, wherein the C.sub.4+ hydrocarbon
feedstock is hydrogenated prior to cracking.
25. The process of claim 1 for cracking C.sub.5+ hydrocarbon
feedstock.
26. The process of claim 1, wherein the hydrocarbon feedstock
comprises saturated paraffins, olefins of general formula
C.sub.nH.sub.2n where n is from 4 to 12 or mixtures thereof.
27. The process of claim 26, wherein the hydrocarbon feedstock
comprises cyclohexane.
28. The process of claim 1, wherein the C.sub.4+ hydrocarbon
feedstock is derived from pyrolysis gasoline or hydrogenated
pyrolysis gasoline.
29. The process of claim 1, wherein the hydrocarbon feedstock
comprises aromatic compounds.
30. The process of claim 1, wherein the process takes place in a
cracker.
31. The process of claim 30, wherein at least a portion of the
cracked hydrocarbon feedstock is recycled back to the cracker and
undergoes further cracking.
32. The process claim 31, wherein the cracked hydrocarbon feedstock
is recycled through a hydrogenator before being returned to the
cracker.
33. The process of claim 1, wherein the temperature is
570-650.degree. C.
34. The process of claim 1, wherein the ratio of Si:Al is 80 to
400.
35. The process of claim 1, wherein the ratio of Si:Al is
85-200.
36. The process of claim 1, wherein the amount of ethylene and
propylene formed is at least 40 wt % carbon.
37. The process of claim 1, wherein the ratio of formed propylene
to ethylene is greater than 1.2:1.
38. The process of claim 1, wherein the amount of propylene formed
is at least 30 wt % carbon.
39. The process of claim 1, wherein at least 60% of the C.sub.4+
hydrocarbon feedstock is cracked.
40. The process of claim 1, wherein the HZSM-5 catalyst is not
hydrothermally treated.
41. The process of claim 1, wherein the HZSM-5 catalyst contains
less than 0.05% wt cations besides than protons.
42. The use of an HZSM-5 catalyst having silicon to aluminum ratio
of 61-1000 as a catalyst in the cracking of C.sub.4+ hydrocarbon
feedstock comprising saturated cyclic hydrocarbons.
Description
[0001] This invention relates to the cracking of higher
hydrocarbons into alkenes using the hydrogen form of pentasil type
zeolite catalysts (HZSM-5). In particular, the invention relates to
the use of HZSM-5 catalysts having specific silicon to aluminium
ratios in the cracking of hydrocarbons to give high yields of
propylene.
[0002] Zeolites are a well-known class of aluminosilicates
containing an (Si, Al).sub.nO.sub.2n framework with negative charge
balanced by cations present in the framework cavities. The nature
of the cations used can vary greatly but typical examples include
potassium, sodium, calcium, silver and hydrogen. The spaces created
within a zeolite framework are relatively large and these so called
"cages" can accommodate large molecules. However, the external
pores of the framework which allow access to the cages of the
zeolite tend to be smaller, allowing control over the sizes of
molecules which can enter and leave the zeolite structure.
[0003] The possibility of size selectivity has made use of zeolite
catalysts an attractive option for the skilled chemist when
carrying out cracking reactions, i.e. reactions where higher
hydrocarbons are broken down into smaller hydrocarbon fragments, as
well as other reactions such as isomerisation reactions or
aromatisation reactions.
[0004] The cracking of higher hydrocarbons into smaller
hydrocarbons, in particular ethylene and propylene, is well known
and has been carried out in the petrochemical industry for many
years. For example, the use of a steam cracker to convert higher
hydrocarbons by pyrolysis into ethylene and propylene is well
known. In general however this reaction produces much more ethylene
than propylene. Zeolite catalysis therefore provides a further
cracking alternative for the skilled chemist.
[0005] U.S. Pat. No. 6,222,087 describes the use of phosphorus
doped ZSM-5 and/or ZSM-1 zeolites in the formation of light olefins
rich in propylene from a hydrocarbon feed containing C.sub.4-7
olefins/paraffins. With a 1-butene feed the invention realises 37%
ethylene and propylene but with a light catalytic naphtha feed only
25% ethylene and propylene is formed.
[0006] U.S. Pat. No. 5,968,342 describes the use of various metal
doped zeolites as catalysts in the conversion of naphtha to
ethylene and propylene. ZSM-5 catalysts in which the proton is
exchanged with silver or copper are suggested to ensure that the
zeolite is substantially free of protons. The text suggests that
HZSM catalysts show poor selectivity for lower olefins. GB 2345294
describes a process similar to that in U.S. Pat. No. 5,968,342,
i.e. one in which a proton free zeolite is employed in a cracking
reaction.
[0007] WO 00/18853 describes ZSM-5 catalysts doped with phosphorus
and a further promoter such as tin or gallium for use in cracking
naphtha to olefins.
[0008] With the increasing demand for propylene throughout the
world, in particular the increasing demand for polypropylene, it
would be useful if more of the heavier petrochemical fractions
could be converted efficiently into propylene for
polymerisation.
[0009] Such heavier petrochemical fractions are often rich in
aromatic compounds notably benzene. The use of aromatic compounds
in fuels is being reduced for environmental reasons resulting in a
potential surplus of aromatic compounds. Moreover, whilst there is
a market for benzene, the market is relatively small. Thus, it is
anticipated that the market for aromatic compounds may become
saturated. It would therefore be useful if aromatic compounds such
as benzene could themselves be converted in high yield into more
marketable olefins such as ethylene and propylene.
[0010] There remains therefore, the need for further cracking
processes to be developed to maximise the yields of ethylene and
propylene which can be obtained from hydrocarbon feeds. In
particular it would be useful if the amounts of propylene produced
could be maximised. It has now been surprisingly found that by
using a hydrogen ZSM-5 catalyst with a particular silicon to
aluminium molar ratio hydrocarbons, such as hydrogenated aromatic
fractions or hydrocarbon fractions containing higher olefins, can
be cracked into ethylene and propylene in very high yields.
Contrary to the teachings of the prior art, these catalysts show
excellent olefin selectivity at the Si/Al molar ratios claimed and
can be used to crack a variety of feeds to give rise to olefins
such as ethylene and propylene in high yield.
[0011] Thus, viewed from one aspect the invention provides a
process for cracking C.sub.4+ hydrocarbons comprising heating the
hydrocarbons at a temperature of 400-800.degree. C. in the presence
of an H-ZSM-5 catalyst having a silicon to aluminium ratio of
61-1000.
[0012] Viewed from another aspect the invention provides the use of
an HZSM-5 catalyst having a silicon to aluminium ratio of 61-1000
as a catalyst in the cracking of C.sub.4+ hydrocarbons.
[0013] The process of the invention may be used on its own or in
combination with other processes such as alternative olefin
producing processes, e.g. pyrolytic steam cracking.
[0014] By C.sub.4+ hydrocarbons is meant that the hydrocarbons
being cracked have at least four carbon atoms. Preferably, the
hydrocarbons being cracked will have between 4 and 20 carbon atoms,
preferably between 5 and 12 carbon atoms, e.g. 6 to 10 carbon
atoms. The hydrocarbon feedstock can be a single pure hydrocarbon
but more usually it will be a mixture of various hydrocarbons, e.g.
a light or heavy naphtha fraction, different condensate fractions
or a hydrogenated pyrolysis gasoline. The hydrocarbons may be
saturated or unsaturated, linear, branched or cyclic and preferably
non-aromatic. Preferred hydrocarbons however will be saturated
paraffins (alkanes such as pentane, hexane, octane, decane,
dodecane), olefins of general formula C.sub.nH.sub.2n where n is
from 4 to 12 or most preferably saturated cyclic hydrocarbons, e.g.
cyclopentane, cyclohexane, methylcyclohexane, dimethylcyclohexane,
methylcyclopentane or decalin. The hydrocarbon feedstock may also
be a mixture of any of the above. In a highly preferred embodiment
the feedstock comprises at least one C4+ olefin, e.g. a C5+ olefin
or diene.
[0015] The hydrocarbon feed may contain aromatic compounds such as
benzene. Such aromatics do not, in general, crack in the presence
of the HZSM-5 catalyst of the invention and are therefore left
unchanged. Aromatic compounds may therefore be used as diluents in
the process of the invention. Preferably however, it would be
useful to convert aromatic compounds into useful olefins.
[0016] Many of the potential feeds to the catalytic cracker contain
high volumes of aromatic compounds, e.g. pyrolysis gasoline or
condensate fractions direct from an oil refinery or an oil
producer.
[0017] In a preferred embodiment therefore, the hydrocarbon feed
may be hydrogenated prior to exposure to the zeolite catalyst. Thus
for example, a hydrocarbon feedstock containing benzene can be
hydrogenated to give cyclohexane prior to cracking. By
hydrogenating the feedstock, the majority of the hydrocarbons fed
to the cracking reactor will be saturated, i.e. alkanes and cyclic
saturated hydrocarbons. This forms a further aspect of the
invention.
[0018] Thus viewed from a further aspect the invention provides a
process comprising hydrogenating a C.sub.4+ hydrocarbon feedstock;
and heating the resulting hydrocarbons at a temperature of
400-800.degree. C. in the presence of an H-ZSM-5 catalyst having a
silicon to aluminium ratio of 61-1000.
[0019] A potential reactor set up could involve diluent free
cracking, e.g. using a hydrogenated pyrolysis gas feed with
recycling of C4+ fractions to the hydrogenator. Alternatively, a
reactor set up could involve the use of an aromatic compound as a
diluent for a hydrocarbon feedstock. Pyrolysis gasoline from a
naphtha steam cracker, which contains a large amount of aromatic
compounds, would, for example, be suitable as such a diluent.
[0020] The cracker may then have a recycling system which passes
through a hydrogenator to convert the diluent into a crackable
hydrocarbon. This may then be fed back into the cracker (along with
fresh diluent) so that in a first pass, the aromatic compound acts
as a diluent prior to hydrogenation and itself being cracked.
Recycling of C4+ fractions could occur simultaneously.
[0021] Preferably the temperature within the cracking reactor
should range from 500-750.degree. C., more preferably
550-700.degree. C., especially 550 to 650.degree. C., e.g.
570.degree. to 650.degree. C. or 570 to 630.degree. C. such as
about 600.degree. C. The residence time of the feed over the
catalyst should be long enough to give substantial conversion of
the feed but not long enough to give rise to the production of
large percentages of aromatic compounds.
[0022] It has been surprisingly found that when the process
described above is employed the yield (calculated as weight of
carbon) of ethylene and propylene can exceed 40 wt % carbon. Thus,
40% by weight of the carbon in the hydrocarbon feedstock is
recovered as ethylene and propylene. Preferably at least 50 wt %
carbon is recovered as ethylene and propylene.
[0023] Whilst the ratio of produced propylene to ethylene can vary,
e.g. more ethylene than propylene or a 1:1 ratio, preferably, the
bulk of the produced product is propylene e.g. at least 25 wt % C,
especially 30 wt % C. It is also preferred if the ratio of produced
propylene to ethylene is greater than 1.2:1, e.g. greater than
1.5:1.
[0024] An important factor in ensuring high percentage yields of
ethylene and propylene is ensuring high conversion of initial
feedstock. It is preferred if at least 50% of the initial feedstock
is cracked, especially at least 60%, e.g. 70%, most especially at
least 80% or at least 90%. At higher temperatures, conversion can
near 100%. The higher the cracking temperature the higher the
conversion. Also at higher temperatures, ethylene selectivity
increases. There is, however, a trade off between temperature and
hence high conversion/desired selectivity and potential for
formation of aromatic compounds. The most preferred temperature to
maximize conversion but minimise benzene formation is in the range
570 to 630.degree. C.
[0025] Viewed from another aspect the invention provides a process
for cracking C.sub.4+ hydrocarbons comprising heating the
hydrocarbons at a temperature of 400-800.degree. C. in the presence
of an H-ZSM-5 catalyst having a silicon to aluminium ratio of
61-1000 and recovering the cracked hydrocarbons;
[0026] wherein at least 40% wt carbon of the cracked C.sub.4+
hydrocarbons is recovered as ethylene and propylene.
[0027] Any non-converted feedstock can of course be recycled back
into the cracker along with any unwanted side products. The
cracking reaction obviously yields hydrocarbons other than ethylene
and propylene. Other products may include ethane, propane, methane,
butane, butene and amounts of C5 and C6+ fractions.
[0028] Catalytic cracking reactors of use in the invention are
known and can operate under the temperatures discussed above using
pressure if necessary, e.g. from 0.1 to 10 atm, preferably 0.3 to 2
atm. The catalytic process can be carried out in, for example, a
fixed bed, moving bed, or fluidised bed reactor and the hydrocarbon
flow can be cocurrent or countercurrent to catalyst flow.
[0029] The catalyst may be formed from fine solid particles having
a size range of from about 0.01 to 10 mm, e.g. 0.2 to 5 mm. Diluent
such as an inert gas (nitrogen), methane or aromatic compounds can
be employed as is known in the art, e.g. to carry the hydrocarbon
gas stream into the reactor. The ratio of inert gas (if used) to
hydrocarbons may range from 0:1 to 1000:1. Careful selection of
diluents may allow further control over the ratio of products
formed. Comprehensive discussions of suitable reactor set up can be
found in the prior art cited above and are known in the art.
[0030] The catalyst employed in the invention, i.e. a HZSM-5
catalyst having a particular Si/Al molar ratio is obtainable from
commercial sources such as Sud-Chemie. Silicon to aluminium ratios
given in the text are molar ratios, i.e. by a Si/Al ratio of 100 is
meant that the molar amount of Si is 100 times the molar amount of
Al. The manufacture of ZSM-5 catalysts is described in U.S. Pat.
No. 3,702,886. Preferred silicon to aluminium ratios are 80-400,
especially 100-300, e.g. 85 to 200. It has been surprisingly found
that at these high ratios, the amounts of formed lower olefins are
high, in contrast to products obtained when using HZSM-5 catalysts
having lower Si/Al ratios. The catalyst is preferably medium pore
(e.g. 10 rings).
[0031] Comprehensive discussions of the manufacture and use of
ZSM-5 catalysts can be found in the prior art cited above. The
amount of catalyst employed will vary depending on the size of the
reactor and the size of the feed but will be readily determined by
the artisan. Catalyst can be continuously added to the cracker if
necessary. It may also be necessary to regenerate the catalyst
using known conditions. It has however, been surprisingly found
that the catalyst of the invention only exhibits slow loss of
activity during the cracking reaction and hence regeneration may be
required only infrequently.
[0032] It is common for the zeolites to be calcined prior to use,
however it is preferred if the HZSM-5 catalyst used in the present
invention is not calcined. Moreover, many zeolite catalysts are
aged prior to use by exposing them to hydrothermal treatment, e.g.
heating at a temperature of 500 to 800.degree. C. in the presence
of steam. The zeolites used in the present invention should not be
heated in this fashion prior to use since this may affect their
light olefin selectivity, i.e. it is not necessary to expose the
zeolite catalysts of the invention to elevated temperatures and
100% steam before use.
[0033] The catalyst of use in the invention is an HZSM-5 species as
hereinbefore described. Hence, the cations within the zeolite
should be protons. There is the possibility however that the
zeolite may be contaminated with other cations, e.g. sodium or
potassium ions. It is preferred if the zeolite of use in this
invention comprises primarily protons with low amounts of other
cations, e.g. exclusively protons and no other cations. For
example, the level of cations (other than protons) should be less
than 0.05% wt, preferably less than 0.01% wt.
[0034] It may be necessary to control the contact time of the
hydrocarbon materials with the catalyst by taking pyrolysis
properties of the hydrocarbons and reaction temperature into
consideration. Contact times, measured as 1/GHSV (gas hourly space
velocity) should be from 0.00002 h to 0.002 h.
[0035] The products of the cracking reaction can be fractionated
using known techniques. Any unwanted products may be recycled to
the reactor so as to increase yield of the desired product(s).
Larger side products may be channeled into a condenser prior to
recycling.
[0036] A suitable reactor set up is now described. The setup used
in the first Example is shown schematically in FIG. 1 and
represents an alternative for carrying out a small scale cracking
reaction. The skilled chemist/chemical engineer can take the
principles described herein for use in an industrial cracking
reactor. Fluid hydrocarbons enter the apparatus via liquid pump (1)
and enter evaporator (3). An inert gas, e.g. nitrogen, is fed to
the evaporator via mass flow controller (2). Gaseous hydrocarbons
formed in the evaporator enter reactor (6) containing catalyst (10)
in oven (5) via flow switch (4). After cracking, the material from
the reactor is transferred to condenser (7) where heavy ends (C9+)
may be collected. In FIG. 1, C1-8 components are fed back through
flow switch (4) and into a mass spectrometer (8) and/or gas
chromatograph (9) for analysis.
[0037] A larger scale apparatus is shown in FIG. 8. The feed is
passed into catalytic naphtha cracker (11) and heated by burning
fuel gas. The cracked feed is transferred to separator (12) where
the components are separated. The separator may also take a feed
from a naphtha steam cracker (13). Fractions can then be isolated
or recycled by to the cracker (11) for further cracking to
occur.
[0038] In FIG. 9, a reactor set up for the cracking of pyrolysis
gasoline is described. To cracker (14) is fed pyrolysis gasoline
and a C4 feed along with hydrogen. The cracked product is
transferred to separator (15) where heavy material and tight
hydrocarbons are separated. The bottoms stream containing the
heavier components is passed through hydrogenator (16) to a further
separator (17) where methane and hydrogen are separated and passed
back to the cracker or isolated. The bottoms stream is also
recycled back to the cracker.
[0039] The top stream from separator (15) is itself passed to a
further separator (19) via pump (18) where C1-3 fractions are
separated from C4/5 fractions. These latter fractions are recycled
to the cracker and the light fractions isolated as the desired
product.
[0040] The invention will now be described with reference to the
following non-limiting examples and Figures. FIG. 1 shows a
potential catalytic cracking reactor set up. FIG. 2 shows the
composition in the gas phase in the reactor effluent from catalyst
A. FIG. 3 shows the composition in the gas phase in the reactor
effluent from catalyst B. FIG. 4 shows the composition in the gas
phase in the reactor effluent from catalyst C. FIG. 5 shows the
composition in the gas phase in the reactor effluent from catalyst
D. FIG. 6 shows conversion of cyclohexane in the presence of HZSM-5
Si/Al=200. FIG. 7 shows selectivity to ethylene and propylene for
HZSM-5 Si/Al=200 fresh and regenerated. FIG. 8 shows a potential
catalytic cracking reactor set up. FIG. 9 shows a potential
catalytic cracking reactor set up. FIG. 10 shows conversion of
cyclohexane over HZSM-5 Si/Al=200 at 400-650.degree. C. using
either pure nitrogen or nitrogen and benzene as diluent. FIG. 11
shows the conversion of C4 and the yield to C2=/C3=and aromates for
catalytic testing of C4 mix over HZSM-5 from Suid Chemie AG
(ZPO31400) with a Si/Al=200. FIG. 12 shows carbon distribution in
the gas effluent after testing the light fraction of pyrolysis
gasoline at increasing temperatures over HZSM-5 from Suid Chemie AG
(ZPO31400) with a Si/Al=200.
EXPERIMENTAL
[0041] The reactor set up depicted in FIG. 1 was employed. The
condenser was maintained at 75.degree. C. and all transfer lines
and switches post evaporator were maintained at a temperature of at
least 60.degree. C.
[0042] The gas chromatograph used was an Agilente micro GC with
four columns. Total analyses of permanent gases and C.sub.1-C.sub.6
took 240 seconds. The gas chromatograph was calibrated with a
C.sub.1-C.sub.4 gas mixture (Standard 1) and a mixture of N.sub.2,
H.sub.2 and C.sub.1-6 (Standard 2). Cyclohexane (CH) was calibrated
by pass analysis. All C.sub.5 compounds were assumed to have the
same calibration factor and all C.sub.6 compounds were assumed to
have the same calibration factor.
Example 1
[0043] Reactions were conducted in the temperature region
400-600.degree. C. with 1 g of catalyst, 50 ml/min N.sub.2 flow
through evaporator and 0.1 ml/min hydrocarbon flow corresponding to
a WHSV [weight hourly space velocity] of 4.7 gCH/(g catalyst*h)
with cyclohexane used as feed. The catalysts were pressed to
tablets and then crushed into particles with particle size 0.2 to
0.5 mm before testing.
[0044] The reactor was first heated to 400.degree. C. under
nitrogen flow before switching to feed from the evaporator. Samples
were taken from the reactor at 4-15 minute intervals. N.sub.2 was
then flushed through the reactor and the reactor heated to
450.degree. C. before feed from the evaporator was again admitted
and the process above repeated. This process was repeated at
500.degree. C., 550.degree. C. and 600.degree. C. before the
reactor was cooled to 400.degree. C. or 450.degree. C. for gas
chromatographic runs to be performed to check for any catalyst
deactivation which may have occurred over the course of the
experiment.
[0045] Four catalysts were examined using the above protocol:
Catalyst A: Kristal 232 ST (Grace Davison)--a standard cracking
catalyst based on rare earth exchanged Y zeolite which has wide
pores (12 rings)
Catalyst B: Valfoor CP 811 BL-25 (PQ) a Beta zeolite with wide
pores (12 rings)
Catalyst C: A HZSM-5 with Si:Al=28 (Sud Chemie) Medium pores (10
ring)
Catalyst D: A HZSM-5 with Si:Al=85 (ZPO 31170) (Sud Chemie) (10
ring)
Catalysts A to C are comparative, catalyst D exemplifies the
invention.
Cracking Catalyst A
[0046] The results for this catalyst are shown in FIG. 2 and in
Appendix Table 1. The catalyst mostly gave high conversion but
deactivated rapidly. Once the temperature returned to 400.degree.
C. after testing at 550.degree. C., activity was almost
non-existent. The carbon mass balance results indicate that a lot
of the carbon forms products that are not analysed in the sampling
system, i.e. coke and C6+ hydrocarbons, The high H.sub.2 yield
points to coke formation. Within the C3 fraction, the majority was
propane as opposed to propene.
Catalyst B
[0047] The results for catalyst B are presented in FIG. 3 and
Appendix table 2. The catalyst gave rise to products in the C5+
range evidenced by the carbon balance results.
Catalyst C
[0048] The results are presented in FIG. 4 and Appendix table 3.
Much C3 product is formed by this catalyst but the fraction is
primarily propane. The conversion is nevertheless high at all
temperatures
Catalyst D
[0049] The results for catalyst D are presented in FIG. 5 and
Appendix table 4. At 600.degree. C. less than 2% of the cyclohexane
is unconverted and there is observed a large selectivity for
propene. The combined ethylene and propylene selectivity among the
gas phase molecules is over 60% at 600.degree. C. The carbon mass
balance at the higher temperature is approximately 80% which means
approximately 50% of the formed product is ethylene and propylene.
The catalyst showed no deactivation on cooling.
Examples 2 to 8
[0050] General Conditions TABLE-US-00001 Amount Flow N.sub.2 flow
of cat. GHSV Temp. Exp. Feed (ml/min) (ml/min) (g).sup.1
(h.sup.-1).sup.7 (.degree. C.) Ex 2 CH 0.1 50 1 4338 400-650 Ex 5
CP 0.1 50 1 4542 400-650 Ex 6 MeCH 0.1 50 1 4131 400-650 Ex 3 CH
0.1 50 1 4338 600 Ex 4 CH 0.1 50 1 4338 600 Ex 7 HPyglf* 0.1 50 1
4286 400-650 Ex 8 HPyg** 0.1 50 1 4110 400-650 *Hydrogenated
pyrolysis gasoline light fraction <125.degree. C. **
Hydrogenated pyrolysis gasoline.
Example 2
[0051] HZSM-5 from Sutd Chemie AG (ZPO31400) with a Si/Al=200 was
pressed into tablets and then crushed into particles with particle
size 0.2-0.5 mm. One gram of catalyst (1 g) particles was tested at
400-650.degree. C. in a quartz fixed bed reactor with on-line GC
analysis. A cyclohexane:N.sub.2 molar ratio of 1:2.2 and a GHSV of
4338 per h was used. Liquid cyclohexane (CH) from Merck (99.6%) was
evaporated at 75.degree. C. into the N.sub.2 stream before entering
the reactor.
[0052] The conversion is calculated on basis of unconverted feed
(nitrogen is used as internal standard) and the carbon selectivity*
in the effluent to propene and ethylene is shown in Table 1.
TABLE-US-00002 TABLE 1 Conversion of CH and C2= and C3= selectivity
over HZSM-5 Temperature Conversion Carbon selectivity* (%)
(.degree. C.) (feed) (%) C.sub.2H.sub.4 C.sub.3H.sub.6 400 17.7 3.6
24.9 500 30.7 9.85 33.1 550 54.6 15.3 37.3 600 77.6 20.8 40.2 650
95.0 29.4 37.8 *Carbon selectivity calculated based on the analysis
of C1-C6 products
Example 3
[0053] The catalyst as described in Example 2 was tested with
cyclohexane using the same experimental procedure as described in
Example 2. The catalyst was tested for 7105 minutes at 600.degree.
C. The conversion as a function of time on stream is shown in FIG.
6. From the results it may easily be derived a half life time of
the catalyst of 3.5 days.
[0054] The carbon selectivity to ethylene and propene (based on
analysis of products C1-C6) is given in FIG. 7.
Example 4
[0055] The catalyst tested in Example 3 was regenerated in 5%
oxygen in He. After regeneration the catalyst was tested with CH
using the same experimental procedure as described in Example 2.
The catalyst was tested for 1397 minutes at 600.degree. C. The
conversion as a function of time is shown in FIG. 6 and
selectivities to ethylene and propylene in FIG. 7.
Example 5
[0056] The catalyst as described in Example 2 was tested with
cyclopentane (CP) using the same experimental procedure as
described in Example 2. A CP:N.sub.2 molar ratio of 1:1.9 and a
GHSV of 4542 h.sup.-1 was used. Liquid CP from Janssen Chimica
16.775.91 (98%) was evaporated at 45.degree. C. into the N.sub.2
stream before entering the reactor.
[0057] The conversion calculated on basis of unconverted feed
(nitrogen is used as internal standard) and the carbon selectivity*
in the effluent to propene and ethylene is shown in Table 2.
TABLE-US-00003 TABLE 2 Conversion of CP and C2= and C3= selectivity
over HZSM-5 Temperature Carbon selectivity* (%) (.degree. C.)
Conversion (feed) (%) C.sub.2H.sub.4 C.sub.3H.sub.6 400 7.4 2.7
20.4 450 20.9 7.0 23.9 500 40.2 12.0 29.4 550 68.7 17.8 33.2 600
90.7 26.5 38.1 650 99.0 35.6 37.0 *Carbon selectivity calculated
based on the analysis of C1-C6 products
Example 6
Methylcyclohexane (MCH) as Feed
[0058] The catalyst as described in Example 2 was tested with MCP
using the same experimental procedure as described in Example 2. A
MCP:N.sub.2 molar ratio of 1:2.7 and a GHSV of 4131 h.sup.-1 was
used. Liquid MCH from Venton 12548 (99%) was evaporated at
90.degree. C. into the N.sub.2 stream before entering the
reactor.
[0059] The conversion calculated on basis of unconverted feed
(nitrogen is used as internal standard) and the carbon selectivity*
in the effluent to propene and ethylene is shown in Table 3.
TABLE-US-00004 TABLE 3 Conversion of MCH and C2= and C3=
selectivity over HZSM-5 Temperature Conversion Carbon selectivity*
(%) (.degree. C.) (feed) (%) C.sub.2H.sub.4 C.sub.3H.sub.6 400 5.0
3.0 20.5 450 12.4 5.3 32.5 500 32.4 8.1 36.6 550 49.5 11.4 39.8 600
72.0 16.1 40.6 650 91.1 23.1 36.9 *Carbon selectivity calculated
based on the analysis of C1-C6 products
Example 7
Cracking of Hydrogenated Pyrolysis Gasoline (Light Fraction)
[0060] Hydrogentated pyrolysis gasoline was obtained from Statoil.
Some of the hydrogenated pyrolysis gas was distilled to boiling
point 125.degree. C. This corresponds to 69.8 wt % of the feed.
98.7 wt % mass balance was obtained during the distillation (this
may indicate that 1.3 wt % of the lightest components have been
lost during the distillation). The total hydrogenated pyrolysis
gasoline (Example 8) and the light fraction <125.degree. C.
(Example 7) were used as feeds.
[0061] Piona analyses of the two feeds were performed and the
results are given in Tables 5 and 7 below. The total feed contains
80.1 wt % naphthenes and the light feed contains 83.2 wt %
naphthenes. The density of the two feeds was measured by weighing
25 ml of the samples. The light fraction <125.degree. C. had a
density of 0.744 g/ml and the total hydrogenated pyrolysis gasoline
had a density of 0.776 g/ml.
[0062] The analysis of the effluent was performed with an Agilente
micro GC with 4 columns. All detectors were TC detectors
Column A: 5 A mol sieve
Run at 30.7 psi and 85.degree. C. with Ar as a carrier gas.
Used for analysis of hydrogen and nitrogen
Column B: Poraplot U
Run at 20.7 psi and 70.degree. C. with He carrier gas.
Used for analysis of methane, ethane and ethylene
Column C: Alumina Plot
Run at 34.3 psi and 115.degree. C. with He carrier gas
Used for analysis of propene, propane and C4's
Column D: OV-1
In order to analysis more of the heavy compounds on the GC, the
temperature of the D column was increased to 150.degree. C.,
compared to 70.degree. C. in earlier analysis. Pressure was 29.6
psi. He carrier gas. This column was used to analyse
C.sub.5-C.sub.9 compounds.
[0063] The catalyst as described in Example 2 was tested with the
light fraction boiling <125.degree. C. of a hydrogenated
pyrolysis gasoline using the experimental conditions as described
in Example 1. A Piona analysis of this feed is given in Table 5
below. The yields to ethene and propene are given Table 6.
TABLE-US-00005 TABLE 5 C3 C4 C5 C6 C7 C8 C9 C10 C11 C12 Paraffins
0.020 0.120 10.202 4.420 1.021 0.165 0.016 0.000 0.000 0.000
n-Paraffins 0.020 0.051 5.857 2.272 0.475 0.082 0.016 0.000 0.000
0.000 iso-Paraffins 0.039 4.345 2.148 0.546 0.083 0.000 0.000 0.000
0.000 Naphthenes 1.967 51.053 20.296 7.988 1.317 0.617 0.000 0.000
Mono-naphthenes 1.967 51.053 20.296 7.988 0.910 0.000 0.000 0.000
Di-naphthenes 0.000 0.408 0.617 0.000 0.000 Aromatics 0.597 0.127
0.000 0.000 0.000 0.000 0.000 Benzenes 0.597 0.127 0.000 0.000
0.000 0.000 0.000 Naphthalene 0.000 Naph-/Olef-benzenes 0.000 0.000
0.000 0.000 0.000 Indenes 0.000 0.000 0.000 0.000 Olefins 0.000
0.000 0.000 0.073 0.000 0.000 0.000 0.000 0.000 n-Olefins 0.000
0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 iso-Olefins 0.000
0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 Naphtheno-olefins
0.000 0.000 0.073 0.000 0.000 0.000 0.000 0.000 Di-olefins 0.000
0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 Other olefins 0.000
0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 Sum 0.020 0.120
12.169 56.071 21.517 8.152 1.333 0.617 0.000 0.000
[0064] TABLE-US-00006 TABLE 6 Conversion of pentanes and
cyclohexane (CH) and yields to ethene and propene. Conv. of
Temperature pentane + i- Conv. of CH Yields (% ) (.degree. C.)
pentane (%) (%) C.sub.2H.sub.4 C.sub.3H.sub.6 400 3.7 1.5 0.1 1.0
450 8.0 9.0 0.5 3.2 500 17.6 30.0 1.8 7.7 550 16.5 49.0 5.5 16.1
600 20.0 72.7 10.5 23.2 650 54.8 93.0 16.8 25.8
[0065] The sum of the C2= and C3=yields are about 43% at
650.degree. C. After running the experiment for approximately 8 h,
the catalyst has been exposed to feed for approximately 115 min.
The conversion at 450.degree. C. was about the same at the end of
the experiment as in the start of the experiment, which points to
no significant deactivation.
Example 8
Hydrogenated Pyrolysis Gasoline (Total Feed)
[0066] The catalyst as described in Example 2 was tested with a
hydrogenated pyrolysis gasoline using the experimental conditions
as described in Example 1. A Piona analysis of this feed is given
in Table 7 below. The yield to ethene and propene are given Table
8. TABLE-US-00007 TABLE 7 Piona analysis of the light fraction from
hydrogenated pyrolysis gasoline Values are given in weight %. C3 C4
C5 C6 C7 C8 C9 C10 C11 C12 Paraffins 0.045 0.165 13.405 4.594 0.747
0.155 0.03 0.096 0.026 0.02 n-Paraffins 0.045 0.111 7.616 1.754
0.345 0.084 0.03 0.019 0.026 0.02 iso-Paraffins 0.053 5.789 2.84
0.402 0.071 0 0.077 0 0 Naphthenes 2.581 36.766 14.657 4.032 8.305
11.689 2.059 0 Mono-naphthenes 2.581 36.766 14.657 3.876 6.484 0.43
0.024 0 Di-naphthenes 0.156 1.82 11.259 2.035 0 Aromatics 0.425
0.081 0.018 0 0 0 0 Benzenes 0.425 0.081 0.018 0 0 0 0 Naphthalene
0 Naph-/Olef-benzenes 0 0 0 0 0 Indenes 0 0 0 0 Olefins 0 0 0 0.051
0 0 0 0.053 0 n-Olefins 0 0 0 0 0 0 0 0 0 iso-Olefins 0 0 0 0 0 0 0
0 0 0 Naphtheno-olefins 0 0 0.051 0 0 0 0 0 Di-olefins 0 0 0 0 0 0
0 0 0 Other olefins 0 0 0 0 0 0 0 0.053 0 Sum 0.045 0.165 15.987
41.785 15.536 4.206 8.335 11.785 2.138 0.02
[0067] TABLE-US-00008 TABLE 8 Conversion of pentanes and
cyclohexane (CH) and yields to ethene and propene. Conv. of
Temperature pentane + i- Conv. of CH Yields (%) (.degree. C.)
pentane (%) (%) C.sub.2H.sub.4 C.sub.3H.sub.6 500 4.0 27.0 1.6 6.3
550 13.0 51.0 5.0 13.4 600 31.5 79.5 9.1 19.5 650 45.0 92.0 15.2
21.6
Example 9
Testing with Aromatic Diluent
[0068] The catalyst described in Example 2 was tested at
400-650.degree. C. in a quartz fixed bed reactor with online GC
analysis. Nitrogen and benzene were used as diluent, along with a
cyclohexane:N.sub.2:benzene molar ratio of 1:1.3:1, and at a GHSV
4356 per h. Liquid 1:1 molar mixture of cyclohexane (CH) from Merck
(99.6%) and benzene from Merck (p.a.) was evaporated at 80.degree.
C. into the nitrogen stream before entering the reactor.
[0069] In FIG. 10 the conversion of CH is compared with the results
from Example 2. In Table 9 the selectivity based on the products
are given. In the selectivity calculations, it is assumed that
benzene is not a product, which influences the results at
temperatures higher than 600.degree. C. TABLE-US-00009 TABLE 9
Conversion and selectivity over HZSM-5 Si/Al = 200 Selectivity (%)
Conversion (%) 400.degree. C. 500.degree. C. 600.degree. C.
Experiment 400.degree. C. 500.degree. C. 600 C.sub.2= C.sub.3=
C.sub.2= C.sub.3= C.sub.2= C.sub.3= Ex 2 17.7 30.7 77.6 3.6 24.9
9.9 33.1 20.8 40.2 Ex 9 19.4 16.0 64.6 1.8 19.4 6.1 32.1 12.9 32.5
*0.1 ml/min CH (liq) + 50 ml/min N2, **0.1 ml/min CH (liq) + 0.098
ml/min Bz (liq) + 27.7 ml/min N2.
Example 10
Testing of a C4 Mix
[0070] HZSM-5 from Sud Chemie AG (ZPO31400) with a Si/Al=200 was
pressed into tablets and then crushed into particles with particle
size 0.2-0.5 mm. One gram of catalyst (1 g) particles was tested at
500-700.degree. C. in a quartz fixed bed reactor with on-line GC
analysis.
[0071] C4 mix obtained from a commercial steam cracker 19 ml/min
and 25 ml/min nitrogen was used as feed. The C4 mix was contained
as a liquid in a container keeping a pressure of 1.5 bar. This was
taken from a stream where the original C4 mix had been exposed to a
partial hydrogenation of butadiene, and thereafter the isobutene
removed by reacting to methyltertbutylether. The analysis of the C4
mix was performed on a separate GC and the analysis is given in
Table 10. TABLE-US-00010 TABLE 10 Analysis of C4 mix obtained from
commercial steam cracker Component Mol % Isobutane 3.3 n-butane
25.9 1-butene 20.2 Trans-butene 35.8 cis-2-butene 15.3
[0072] The conversion of the C4's and the yield to C2=/C3= and
aromates are given in FIG. 11. This example shows that the process
is well suited for cracking of C4 mix into light olefins and up to
50% yield to C2=/C3= is achieved at 600-650.degree. C.
Example 11
Using Light Fraction of Pyrolysis Gasoline as Feed
[0073] The catalyst (1 g) as described in Example 2 was tested with
the light fraction of pyrolysis gasoline as feed. The pyrolysis
gasoline was distilled and the fraction boiling at
.ltoreq.125.degree. C. was used as feed. This corresponds to 82% of
the pyrolysis gasoline sample. Analysis of the light fraction of
the pyrolysis gasoline was performed on a separate GC and the
analysis is given in Table 11. This light fraction had a density of
0.798 g/ml.
[0074] The liquid feed was fed with 0.1 ml/min and evaporated into
a 50 ml/min nitrogen stream before entering the reactor.
TABLE-US-00011 TABLE 11 Analysis of "light fraction" of pyrolysis
gasoline (not hydrogenated) Component Wt % Benzene 35.5 Toluene
10.9 Other C.sub.5-C.sub.8, mainly C.sub.5 49.0 Xylenes +
ethylbenzene 4.6
[0075] The carbon distribution in the gas effluent is given in FIG.
12. This example shows that the process cracks 20% of the light
naphta into C2=/C3=. This corresponds to 16.4% of the total
pyrolysis gasoline. TABLE-US-00012 TABLE 1 The cracking catalyst A
Mol % i gas phase sum mol % in C sum sum sum sum Temp loop balance
H.sub.2 CH.sub.4 C.sub.2H.sub.4 C.sub.2H.sub.6 C.sub.3H.sub.6
C.sub.3H.sub.8 C.sub.4 C.sub.5 C.sub.6prod unconv C*n 400 102.9
80.6 0.1 0.0 0.1 0.0 0.0 0.5 0.7 0.2 3.8 22.0 28.5 400 103.2 83.8
0.1 0.0 0.1 0.0 0.0 0.4 0.5 0.7 3.5 23.0 27.8 450 101.5 77.4 0.4
0.0 0.4 0.0 0.3 2.0 1.8 1.1 5.4 16.3 52.9 450 103.5 105.5 0.5 0.2
0.4 0.1 0.3 2.1 2.0 1.3 7.0 21.0 65.1 500 101.3 75.8 2.6 0.8 1.0
0.3 1.3 4.9 3.2 1.3 5.7 10.7 75.2 500 104.2 130.9 2.6 0.7 0.8 0.2
1.2 4.3 3.2 1.0 9.0 20.0 91.1 550 101.6 74.4 8.9 2.3 1.3 0.7 2.4
4.9 2.5 1.0 5.0 8.7 73.0 550 123.5 88.3 8.6 1.8 1.1 0.5 2.3 4.0 2.3
1.1 6.3 17.6 76.2 550 103.3 100.1 6.8 1.4 0.8 0.4 1.9 2.8 1.6 0.9
4.8 18.3 58.0 550 209.7 82.5 6.4 1.3 0.7 0.3 1.9 2.3 1.4 0.8 4.1
18.0 49.9 600 103.8 89.2 7.6 1.2 0.7 0.2 2.6 1.0 0.9 0.6 2.6 20.1
35.7 600 113.9 82.4 6.5 0.9 0.5 0.2 2.1 0.7 0.8 0.6 2.3 22.6 30.8
Product carbon % in gas phase sum sum sum Temp Metan C.sub.2H.sub.4
C.sub.2H.sub.6 C.sub.3H.sub.6 C.sub.3H.sub.8 C.sub.4 C.sub.5
C.sub.6p 400 0.0 0.7 0.1 0.5 5.3 10.0 3.7 79.8 400 0.0 0.6 0.1 0.5
4.1 7.6 11.9 75.3 450 0.0 1.4 0.2 1.7 11.4 14.0 10.0 61.5 450 0.3
1.2 0.2 1.5 9.7 12.5 10.1 64.5 500 1.1 2.6 0.7 5.1 19.5 17.1 8.6
45.2 500 0.8 1.9 0.5 3.8 14.1 14.2 5.6 59.0 550 3.2 3.6 1.9 10.0
20.2 13.7 6.6 40.7 550 2.4 2.9 1.4 8.9 15.8 12.2 7.1 49.4 550 2.5
2.9 1.2 9.9 14.5 11.3 7.5 50.1 550 2.6 3.0 1.2 11.3 14.0 11.2 7.5
49.2 600 3.4 3.6 1.3 21.5 8.2 9.9 8.5 43.6 600 2.9 3.4 1.1 20.8 7.1
11.0 9.0 44.7
[0076] TABLE-US-00013 TABLE 2 Catalyst B Zeolite Beta Mol % in gas
phase sum C Temp mol % in balance sum sum sum sum (.degree. C.)
loop (%) H.sub.2 CH.sub.4 C.sub.2H.sub.4 C.sub.2H.sub.6
C.sub.3H.sub.6 C.sub.3H.sub.8 C.sub.4 C.sub.5 C.sub.6prod unconv
C*n 400 98.7 71.7 5.3 0.8 0.5 0.1 0.3 4.4 7.0 4.1 6.9 3.1 106.3 400
102.0 79.7 5.3 0.6 0.4 0.0 0.3 2.4 4.2 2.8 9.2 8.1 95.8 400 99.8
55.7 3.3 0.4 0.3 0.0 0.2 1.5 2.2 1.8 5.9 8.7 58.9 400 98.6 45.1 2.6
0.3 0.2 0.0 0.2 1.0 1.7 1.2 4.8 8.0 45.8 450 97.9 42.6 9.6 1.0 0.9
0.1 0.6 3.4 3.6 1.8 2.8 4.0 55.0 450 96.0 34.6 7.5 0.8 0.6 0.1 0.5
2.0 2.2 1.1 2.2 5.1 36.9 450 96.1 30.3 6.4 0.7 0.5 0.1 0.4 1.5 1.7
0.9 1.9 5.2 30.2 500 94.1 30.0 14.2 1.7 1.2 0.3 1.0 3.1 2.6 0.8 1.5
2.0 40.5 500 94.6 30.1 13.0 1.6 1.1 0.2 1.0 2.6 2.3 0.8 1.6 2.7
37.8 550 93.2 27.5 21.2 2.8 1.6 0.5 1.2 2.6 1.4 0.2 2.6 0.4 40.1
550 93.8 29.9 18.9 2.5 1.5 0.4 1.5 2.2 1.6 0.3 2.5 1.3 40.6 450
97.6 45.2 5.1 0.5 0.3 0.0 0.3 0.4 0.6 0.4 1.5 12.3 17.1 Product
carbon % in gas phase Temp sum sum sum (.degree. C.) Metan
C.sub.2H.sub.4 C.sub.2H.sub.6 C.sub.3H.sub.6 C.sub.3H.sub.8 C.sub.4
C.sub.5 C.sub.6p 400 0.8 0.9 0.1 0.8 12.6 26.5 19.5 38.8 400 0.6
0.8 0.1 1.1 7.5 17.6 14.6 57.7 400 0.7 0.9 0.1 0.9 7.4 14.7 15.0
60.3 400 0.7 1.0 0.1 1.6 6.8 14.6 12.7 62.5 450 1.9 3.1 0.4 3.5
18.4 26.1 16.3 30.3 450 2.2 3.2 0.4 3.9 16.0 23.6 15.0 35.6 450 2.3
3.3 0.4 4.3 15.1 22.8 14.2 37.6 500 4.3 6.1 1.3 7.3 22.8 25.9 9.8
22.5 500 4.1 6.0 1.3 7.9 20.7 24.8 10.1 25.0 550 7.1 7.7 2.4 8.7
19.4 14.1 2.4 38.2 550 6.0 7.3 2.0 10.9 16.1 16.0 4.1 37.4 450 2.9
3.1 0.3 5.6 7.3 15.1 11.4 54.4
[0077] TABLE-US-00014 TABLE 3 Catalyst C, HZSM-5 with Si/Al = 28
Mol % in gas phase sum C Temp mol % in balance sum sum sum sum
(.degree. C.) loop (%) H.sub.2 CH.sub.4 C.sub.2H.sub.4
C.sub.2H.sub.6 C.sub.3H.sub.6 C.sub.3H.sub.8 C.sub.4 C.sub.5
C.sub.6prod unconv C*n 400 101.5 48.0 1.2 0.2 0.7 0.3 0.9 15.4 4.3
0.8 0.6 2.9 76.4 400 101.0 51.4 1.2 0.2 0.8 0.3 1.0 16.6 4.6 0.9
0.7 2.6 82.1 400 100.6 47.8 1.1 0.2 0.8 0.3 0.9 15.7 4.2 0.9 0.7
2.5 77.5 400 100.1 46.9 1.1 0.2 0.7 0.3 0.9 15.2 4.3 0.8 0.7 2.5
75.9 400 100.3 48.4 1.1 0.2 0.7 0.3 0.9 15.7 4.4 0.8 0.7 2.5 77.9
400 99.9 49.7 1.1 0.2 0.8 0.3 1.0 15.8 4.4 0.9 0.7 2.7 78.6 400
100.3 48.5 1.2 0.2 0.8 0.3 0.9 15.7 4.4 0.9 0.6 2.5 78.0 400 101.8
50.8 1.3 0.2 0.8 0.3 1.0 16.1 4.5 0.9 0.6 3.0 79.8 400 100.2 43.6
1.2 0.2 0.7 0.3 0.9 14.2 3.9 0.8 0.6 2.5 70.8 400 100.9 44.4 1.2
0.2 0.7 0.3 0.9 14.2 3.9 0.8 0.6 2.9 70.9 450 99.5 40.4 3.0 0.6 1.5
0.6 1.8 14.6 3.6 0.5 0.7 0.4 74.8 450 100.0 40.1 2.9 0.6 1.5 0.6
1.7 14.7 3.6 0.5 0.7 0.3 75.2 450 100.4 44.3 3.2 0.6 1.5 0.7 1.8
15.9 3.9 0.6 0.8 0.3 81.2 500 99.6 40.6 6.3 1.9 2.8 1.2 2.6 12.5
2.6 0.2 1.0 0.1 72.9 500 98.9 37.2 6.2 1.7 2.7 1.1 2.4 11.5 2.4 0.2
1.0 0.0 68.2 550 94.7 27.5 10.3 0.0 4.2 1.6 2.3 6.1 0.7 0.0 1.5 0.1
49.0 Product carbon % in gas phase Temp sum sum sum (.degree. C.)
Metan C.sub.2H.sub.4 C.sub.2H.sub.6 C.sub.3H.sub.6 C.sub.3H.sub.8
C.sub.4 C.sub.5 C.sub.6p 400 0.3 2.0 0.8 3.7 60.4 22.6 5.4 4.9 400
0.2 1.9 0.8 3.5 60.7 22.6 5.4 4.8 400 0.2 2.0 0.8 3.7 60.7 21.8 5.7
5.1 400 0.2 1.9 0.8 3.7 60.3 22.4 5.3 5.3 400 0.2 1.9 0.8 3.6 60.3
22.6 5.4 5.1 400 0.2 1.9 0.8 3.7 60.2 22.4 5.5 5.2 400 0.2 1.9 0.8
3.6 60.5 22.3 5.6 5.0 400 0.3 2.0 0.8 3.6 60.4 22.6 5.5 4.9 400 0.3
2.1 0.8 3.9 60.2 22.3 5.4 5.0 400 0.3 2.1 0.8 3.9 60.0 22.1 5.8 5.0
450 0.8 4.0 1.6 7.1 58.5 19.0 3.3 5.8 450 0.7 3.9 1.6 7.0 58.5 19.1
3.4 5.7 450 0.7 3.8 1.6 6.8 58.8 19.0 3.6 5.6 500 2.6 7.8 3.4 10.5
51.3 14.2 1.5 8.6 500 2.5 8.0 3.4 10.8 50.7 14.2 1.5 9.0 550 0.0
17.1 6.3 14.3 37.1 5.9 0.4 18.9
[0078] TABLE-US-00015 TABLE 4 Catalyst D, HZSM-5 with Si/Al = 85
Mol % in gas phase sum C Temp mol % in balance sum sum sum sum
(.degree. C.) loop (%) H.sub.2 CH.sub.4 C.sub.2H.sub.4
C.sub.2H.sub.6 C.sub.3H.sub.6 C.sub.3H.sub.8 C.sub.4 C.sub.5
C.sub.6prod unconv C*n 600 103.5 86.0 9.2 2.2 16.1 2.0 11.5 4.2 4.3
1.1 1.0 0.3 114.8 600 97.7 81.9 8.9 1.9 13.8 1.0 11.6 4.2 4.3 0.7
1.0 0.3 106.1 600 96.7 82.2 8.8 1.8 13.5 1.0 11.6 4.3 4.3 0.7 1.0
0.3 105.3 600 96.0 79.3 8.7 1.7 13.1 0.9 11.4 4.1 4.2 0.7 1.0 0.3
102.8 600 96.2 81.9 8.7 2.7 13.2 1.0 13.6 2.3 4.3 0.7 1.0 0.3 104.9
600 96.4 79.4 8.6 1.7 13.1 1.0 11.4 4.1 4.2 0.7 1.0 0.3 103.3 550
97.0 80.9 6.7 0.8 8.8 0.6 9.9 6.0 5.3 1.1 0.9 2.7 99.3 550 96.9
78.9 6.5 0.8 8.8 0.6 9.8 5.9 5.2 1.1 0.9 2.7 98.0 550 96.5 80.7 6.5
0.8 8.8 0.6 9.9 6.0 5.2 1.1 0.9 2.7 99.0 550 96.9 81.9 6.5 0.8 8.8
0.6 10.0 6.1 5.3 1.1 0.9 2.7 100.3 550 102.1 81.2 6.8 1.1 9.9 0.7
10.0 6.2 5.4 1.1 0.9 3.0 104.0 550 98.0 80.7 6.5 0.9 9.0 0.7 9.9
6.0 5.3 1.1 0.9 2.8 100.0 500 105.5 101.8 3.6 0.7 5.4 0.4 6.7 7.1
5.9 3.0 1.6 10.3 102.1 450 101.6 90.5 1.3 0.5 2.0 0.1 2.6 4.6 2.8
0.7 0.2 19.6 42.7 400 100.4 96.2 0.3 0.4 0.5 0.0 0.7 1.9 1.0 0.3
0.6 26.2 17.6 400 100.7 103.4 0.3 0.2 0.5 0.0 0.7 2.0 1.0 0.3 0.6
27.6 18.5 Product carbon % in gas phase Temp sum sum sum (.degree.
C.) Metan C.sub.2H.sub.4 C.sub.2H.sub.6 C.sub.3H.sub.6
C.sub.3H.sub.8 C.sub.4 C.sub.5 C.sub.6p 600 1.9 28.0 3.5 30.2 11.1
15.0 4.8 5.4 600 1.8 26.0 1.8 32.8 11.9 16.3 3.4 5.9 600 1.7 25.6
1.8 33.0 12.1 16.5 3.3 5.9 600 1.7 25.5 1.8 33.2 12.0 16.4 3.4 6.0
600 1.6 25.1 1.8 38.8 6.7 16.6 3.5 5.9 600 1.7 25.4 1.8 33.2 12.0
16.4 3.4 6.0 550 0.8 17.7 1.2 29.9 18.1 21.1 5.6 5.5 550 0.8 17.9
1.2 29.9 18.0 21.0 5.6 5.6 550 0.8 17.8 1.2 29.9 18.1 21.1 5.6 5.4
550 0.8 17.5 1.2 30.0 18.2 21.2 5.6 5.5 550 1.1 19.1 1.4 28.9 17.9
20.7 5.5 5.3 550 0.9 18.0 1.3 29.7 18.0 21.1 5.6 5.4 500 0.8 11.7
0.9 21.6 22.9 25.1 4.9 9.5 450 1.2 9.3 0.6 18.2 32.2 26.7 8.3 3.4
400 2.0 5.3 0.3 11.6 32.1 21.6 7.7 19.4 400 1.3 5.2 0.3 11.1 31.9
21.5 7.7 21.0
* * * * *